U.S. patent number 5,401,385 [Application Number 08/104,835] was granted by the patent office on 1995-03-28 for selective upgrading of naphtha.
This patent grant is currently assigned to UOP. Invention is credited to Paula L. Bogdan, R. Joe Lawson, J. W. Adriaan Sachtler, Robert J. Schmidt.
United States Patent |
5,401,385 |
Schmidt , et al. |
March 28, 1995 |
Selective upgrading of naphtha
Abstract
A process combination is disclosed to selectively upgrade
catalytically cracked gasoline to obtain products suitable for
further upgrading to reformulated fuels. A naphtha feedstock,
preferably heavy naphtha, is hydrogenated to saturate aromatics,
followed by selective isoparaffin synthesis to yield light and
heavy synthesis naphtha and isobutane. The heavy synthesis naphtha
may be processed by reforming, light naphtha may be isomerized, and
isobutane may be upgraded by dehydrogenation, etherification and/or
alkylation to yield gasoline components from the process
combination suitable for production of reformulated gasoline.
Inventors: |
Schmidt; Robert J. (Rolling
Meadows, IL), Bogdan; Paula L. (Mount Prospect, IL),
Lawson; R. Joe (Arlington Heights, IL), Sachtler; J. W.
Adriaan (Des Plaines, IL) |
Assignee: |
UOP (Des Plaines, IL)
|
Family
ID: |
27121634 |
Appl.
No.: |
08/104,835 |
Filed: |
August 10, 1993 |
Related U.S. Patent Documents
|
|
|
|
|
|
|
Application
Number |
Filing Date |
Patent Number |
Issue Date |
|
|
795573 |
Nov 21, 1991 |
5242576 |
|
|
|
796562 |
Nov 21, 1991 |
5235120 |
|
|
|
Current U.S.
Class: |
208/57; 208/78;
208/79; 208/80; 208/89 |
Current CPC
Class: |
C10G
59/02 (20130101); C10G 59/06 (20130101); C10G
65/08 (20130101); C10G 69/08 (20130101); C10L
1/023 (20130101); C10G 2400/02 (20130101) |
Current International
Class: |
C10G
65/00 (20060101); C10G 59/00 (20060101); C10G
65/08 (20060101); C10G 69/08 (20060101); C10L
1/00 (20060101); C10G 59/02 (20060101); C10L
1/02 (20060101); C10G 59/06 (20060101); C10G
69/00 (20060101); C10G 045/00 (); C10G 059/02 ();
C10G 059/06 () |
Field of
Search: |
;208/57,79,80,89,78 |
References Cited
[Referenced By]
U.S. Patent Documents
Primary Examiner: Sneed; Helen M. S.
Assistant Examiner: Griffin; Walter D.
Attorney, Agent or Firm: McBride; Thomas K. Spears, Jr.;
John F. Conser; Richard E.
Parent Case Text
Cross-Reference to Related Applications
This application is a continuation-in-part of prior application
Ser. No. 795,573, filed Nov. 21, 1991, now U.S. Pat. No. 5,242,576
and also a continuation-in-part of prior application Ser. No.
796,562, filed Nov. 21, 1991, now U.S. Pat. No. 5,235,120 both of
which are incorporated herein by reference.
Claims
WE CLAIM AS OUR INVENTION
1. A process combination for selectively upgrading a catalytically
cracked gasoline feedstock having a final boiling point of from
about 160.degree. to 230.degree. C. to obtain lower-boiling
hydrocarbons comprising superequilibrium isobutane, comprising the
steps of:
(a) contacting the gasoline feedstock in a hydrogenation zone with
a hydrogenation catalyst in the presence of hydrogen at a pressure
of from about 10 to 100 atmospheres, a temperature of at least
30.degree. C. and a liquid hourly space velocity of from about 1 to
8 to produce a saturated intermediate; and,
(b) converting the saturated intermediate in a
selective-isoparaffin-synthesis zone maintained at a pressure of
from about 10 to 100 atmospheres, a temperature of between about
50.degree. C. and 350.degree. C. and a liquid hourly space velocity
of between about 0.5 and 20 with a solid acid selective
isoparaffin-synthesis catalyst in the presence of hydrogen and
recovering synthesis product containing at least 8 volume % butanes
and having a reduced final boiling point relative to the gasoline
feedstock.
2. The process combination of claim 1 wherein the saturated
intermediate is transferred from the hydrogenation zone to the
selective-isoparaffin-synthesis zone without separation of hydrogen
or light hydrocarbons.
3. The process combination of claim 2 wherein the saturated
intermediate is transferred from the hydrogenation zone to the
selective-isoparaffin-synthesis zone without heating.
4. The process combination of claim 1 wherein the hydrogenation
catalyst comprises a supported platinum-group metal component.
5. The process combination of claim 4 wherein the platinum-group
metal component comprises a platinum component.
6. The process combination of claim 4 wherein the
hydrogenation-catalyst support comprises a refractory inorganic
oxide.
7. The process combination of claim 4 wherein the hydrogenation
catalyst further comprises one or more metals of Group VIB (6),
Group VIII (8-10) and Group IVA (14).
8. The process combination of claim 1 further comprising recovering
a stream containing superequilibrium isobutane from the
selective-isoparaffin-synthesis zone of step (b).
9. The process combination of claim 1 wherein the selective
isoparaffin-synthesis catalyst comprises a platinum-group metal
component on a chlorided inorganic-oxide support.
10. The process combination of claim 9 wherein the platinum-group
metal component comprises a platinum component.
11. The process combination of claim 9 wherein the inorganic-oxide
support comprises alumina.
12. The process combination of claim 9 wherein the catalyst
comprises a Friedel-Crafts metal halide.
13. The process combination of claim 12 wherein the Friedel-Crafts
metal halide comprises aluminum chloride.
14. The process combination of claim 1 wherein the feedstock
comprises straight-run naphtha in admixture with the catalytically
cracked gasoline.
15. The process combination of claim 8 wherein the synthesis
product of step (b) is separated to obtain a light synthesis
naphtha comprising pentanes and a heavy synthesis naphtha
comprising C.sub.7 and C.sub.8 hydrocarbons.
16. A process combination for selectively upgrading a catalytically
cracked gasoline feedstock having a final boiling point of from
about 160.degree. to 230.degree. C. to obtain lower-boiling
hydrocarbons comprising superequilibrium isobutane, comprising the
steps of:
(a) separating the gasoline feedstock to obtain a heart-cut naphtha
fraction comprising C.sub.7 and C.sub.8 hydrocarbons and a heavy
naphtha fraction comprising C.sub.10 hydrocarbons;
(b) contacting the heavy naphtha fraction in a hydrogenation zone
with a hydrogenation catalyst in the presence of hydrogen at a
pressure of from about 10 to 100 atmospheres, a temperature of at
least 30.degree. C. and a liquid hourly space velocity of from
about 1 to 8 to produce a saturated intermediate; and,
(c) converting the saturated intermediate in a
selective-isoparaffin-synthesis zone maintained at a pressure of
from about 10 to 100 atmospheres, a temperature of between about
50.degree. C. and 350.degree. C. and a liquid hourly space velocity
of between about 0.5 and 20 with a solid acid selective
isoparaffin-synthesis catalyst in the presence of hydrogen and
recovering a synthesis product containing at least 8 volume %
butanes and having a reduced final boiling point relative to the
gasoline feedstock.
17. A process combination for selectively upgrading a contaminated
catalytically cracked gasoline feedstock having a final boiling
point of from about 160.degree. to 230.degree. C. to obtain
lower-boiling hydrocarbons comprising superequilibrium isobutane,
comprising the steps of:
(a) hydrotreating the gasoline feedstock to convert sulfurous and
nitrogenous compounds and obtain a naphtha feedstock;
(b) separating the naphtha feedstock to obtain a heart-cut naphtha
fraction comprising C.sub.7 and C.sub.8 hydrocarbons and a heavy
naphtha fraction comprising C.sub.10 hydrocarbons;
(c) contacting the heart-cut naphtha fraction with a reforming
catalyst, comprising a supported platinum-group metal component, in
a catalytic-reforming zone maintained at a pressure of from about
atmospheric to 20 atmospheres, a temperature of from 260.degree. to
560.degree. C. and a liquid hourly space velocity of from about 1
to 40 and recovering a stabilized reformate;
(d) contacting the heavy naphtha fraction in a hydrogenation zone
with a hydrogenation catalyst in the presence of hydrogen at a
pressure of from about 10 to 100 atmospheres, a temperature of at
least 30.degree. C. and a liquid hourly space velocity of from
about 1 to 8 to produce a saturated intermediate;
(e) converting the saturated intermediate in a
selective-isoparaffin-synthesis zone maintained at a pressure of
from about 10 to 100 atmospheres, a temperature of between about
50.degree. and 350.degree. C. and a liquid hourly space velocity of
between about 0.5 and 20 with a solid acid selective
isoparaffin-synthesis catalyst in the presence of hydrogen and
recovering a synthesis product containing at least 8 volume %
butanes and having a reduced final boiling point relative to the
gasoline feedstock.
18. The process combination of claim 17 wherein the synthesis
product of step (e) is separated to obtain a light synthesis
naphtha comprising pentanes and a heavy synthesis naphtha
comprising C.sub.7 and C.sub.8 hydrocarbons.
19. The process combination of claim 18 wherein the heavy synthesis
naphtha is contacted with a catalyst, comprising a supported
platinum-group metal component, in a reforming zone to obtain a
reformed synthesis product.
Description
BACKGROUND OF THE INVENTION
1. Field of the Invention
This invention relates to an improved process combination for the
conversion of hydrocarbons, and more specifically for the selective
upgrading of naphtha fractions by a combination of aromatics
removal and selective isoparaffin synthesis.
2. General Background
The widespread removal of lead antiknock additive from gasoline and
the rising fuel-quality demands of high-performance
internal-combustion engines have compelled petroleum refiners to
install new and modified processes for increased "octane," or knock
resistance, in the gasoline pool. Refiners have relied on a variety
of options to upgrade the gasoline pool, including higher-severity
catalytic reforming, higher FCC (fluid catalytic cracking) gasoline
octane, isomerization of light naphtha and the use of oxygenated
compounds. Such key options as increased reforming severity and
higher FCC gasoline octane result in a higher aromatics content of
the gasoline pool, through the production of high-octane aromatics
at the expense of low-octane heavy paraffins. Current gasolines
generally have aromatics contents of about 30% or higher, and may
contain more than 40% aromatics.
Currently, refiners are faced with the prospect of supplying
reformulated gasoline to meet tightened automotive emission
standards. Reformulated gasoline would differ from the existing
product in having a lower vapor pressure, lower final boiling
point, increased content of oxygenates, and lower content of
olefins, benzene and aromatics. The oxygen content of gasoline will
be 2 mass% or more in many areas. Gasoline aromatics content is
likely to be lowered into the 20-25% range in major urban areas,
and low-emission gasoline containing less than 15 volume% aromatics
is being advocated for some areas with severe pollution problems.
Distillation end points also could be lowered, further restricting
aromatics content since the high-boiling portion of the gasoline
which thereby would be eliminated usually is an aromatics
concentrate. End point often is characterized as the 90%
distillation temperature, currently limited to a maximum of
190.degree. C. and averaging 165.degree.-170.degree. C., which
could be reduced to around 150.degree. C. in some cases.
Since aromatics have been the principal source of increased
gasoline octanes during the recent lead-reduction program, severe
restriction of the aromatics content and high-boiling portion will
present refiners with processing problems. Currently applicable
technology includes such processes as recycle isomerization of
light naphtha and generation of additional light olefins through
fluid catalytic cracking and isobutane through isomerization as
feedstock to an alkylation unit. In,creased blending of oxygenates
such as methyl tertiary-butyl ether (MTBE) and ethanol will be an
essential part of the reformulated-gasoline program, but feedstock
supplies will become stretched. Selective isoparaffin synthesis to
produce desirable gasoline components is known but has not yet
become attractive for commercialization.
A process designated as "1-cracking" for increasing the yield of
naphtha and isobutane is disclosed in U.S. Pat. No. 3,692,666
(Pollitzer). U.S. Pat. No. 3,788,975 (Donaldson) teaches a
combination process for the production and utilization of aromatics
and isobutane. The combination includes selective, production of
isobutane from naphtha followed by a combination of processes
including catalytic reforming, aromatic separation, alkylation,
isomerization, and alehydrogenation to yield alkylation feedstock.
The paraffinic stream from aromatic extraction is returned to the
l-cracking step. Neither Pollitzer nor Donaldson disclose the
present process combination, however, nor do they recognize any
problem from processing aromatics-containing charge stocks.
A combination process for gasoline production is disclosed in U.S.
Pat. No. 3,933,619 (Kozlowski). High-octane, low-lead or unleaded
gasoline is produced by hydrocracking a hydrocarbon feedstock to
obtain butane, pentane-hexane, and C.sub.7 + hydrocarbons, and the
C.sub.7 + fraction may be sent to a reformer. U.S. Pat. No.
4,647,368 (McGuiness et al.) discloses a method for upgrading
naphtha by hydrotreating, hydrocracking over zeolite beta,
recovering isobutane, C.sub.5 -C.sub.7 isoparaffins and a higher
boiling stream, and reforming the latter stream. U.S. Pat. No.
4,897,177 (Nadler) discloses separation of naphtha with light and
heavy fractions being reformed, respectively, over monofunctional
and bifunctional catalysts. These references do not teach or
suggest a process combination including selective isoparaffin
synthesis, however.
Isomerization of C.sub.4 -C.sub.6 paraffins with a hydrogenation
zone upstream to saturate benzene is taught in U.S. Pat. No.
5,003,118 (Low et al.). U.S. Pat. No. 2,493,499 (Perry) teaches
saturation of aromatics and olefins by hydrogenation prior to
isomerization. Both of these references teach that hydrogenation
reduces subsequent cracking, in contrast to the context of the
selective isoparaffin synthesis process of the present
invention.
The prior art, therefore, contains elements of the present
invention. There is no suggestion to combine the elements, however,
nor of the surprising benefits in selectivity that accrue from the
present process combination to obtain intermediate hydrocarbons
suitable for producing reformulated gasoline.
SUMMARY OF THE INVENTION
It is an object of the present invention to provide an improved
process combination to upgrade-naphtha to gasoline. A specific
object is to improve selectivity in producing hydrocarbons suitable
for producing reformulated gasoline.
This invention is based on the discovery that a process combination
comprising aromatics hydrogenation followed by selective
isoparaffin synthesis provides a more even temperature profile
during the synthesis step along with surprising improvements in the
isoparaffin content of the synthesis product.
A broad embodiment of the present invention is directed to a
process combination comprising hydrogenation of aromatics in a
naphtha feedstock followed by selective isoparaffin synthesis from
the hydrogenated naphtha to yield a synthesis product comprising
isobutane and synthesis naphtha with reduced end point. Preferably,
the hydrogenation and selective isoparaffin synthesis are
accomplished in the same hydrogen circuit and the heat of
hydrogenation raises the temperature of the saturated intermediate
to that required for the selective isoparaffin synthesis.
Optionally, heavy synthesis naphtha is separated from the products
and reformed and light naphtha and isobutane are upgraded to useful
gasoline blending components.
One feedstock embodiment is the upgrading of catalytically cracked
gasoline by selective isoparaffin synthesis.
Optionally, a naphtha feedstock is separated into selected naphtha
fractions, with selective isoparaffin synthesis from heavy naphtha
to yield a product comprising isobutane and synthesis naphtha with
reduced end point. A heart-cut naphtha from the separation step may
be reformed, optionally in combination with a heavy synthesis
product.
These as well as other objects and embodiments will become apparent
from the detailed description of the invention.
BRIEF DESCRIPTION OF THE DRAWINGS
FIG. 1 represents a simplified block flow diagram showing the
arrangement of the major sections of the present invention.
FIG. 2 shows reactor temperature profiles when using the process
combination of the invention in comparison to those of the prior
art to process naphtha feedstock derived from a paraffinic crude
oil.
FIG. 3 compares product butane and pentane isomer ratios as well as
conversion for processes of the invention and prior art
corresponding to the cases of FIG. 1.
FIG. 4 shows reactor temperature profiles when using the process
combination of the invention in comparison to those of the prior
art to process feedstock derived from a catalytically cracked
gasoline.
DESCRIPTION OF THE PREFERRED EMBODIMENTS
To reiterate, a broad embodiment of the present invention is
directed to a process combination comprising hydrogenation of
aromatics in a naphtha feedstock followed by selective isoparaffin
synthesis from the hydrogenated naphtha to yield a product
comprising isobutane and synthesis naphtha with reduced end point.
Usually the process combination is integrated into a petroleum
refinery comprising crude-oil distillation, reforming, cracking and
other processes known in the art to produce finished gasoline and
other petroleum products.
The naphtha feedstock to the present process combination will
comprise paraffins, naphthenes, and aromatics, and may comprise
small amounts of olefins, boiling within the gasoline range.
Feedstocks which may be utilized include straight-run naphthas,
natural gasoline, synthetic naphthas, thermal gasoline,
catalytically cracked gasoline, partially reformed naphthas or
raffinates from extraction of aromatics. The distillation range
generally is that of a full-range naphtha, having an initial
boiling point typically from 0.degree. to 100.degree. C. and a
final boiling point of from about 160.degree. to 230.degree. C.;
more usually, the initial boiling range is from about 40.degree. to
80.degree. C. and the final boiling point from about 1750.degree.
to 200.degree. C. Generally the naphtha feedstock contains less
than about 30 mass % C.sub.6 and lighter hydrocarbons, and usually
less than about 20 mass % C.sub.6 -, since the objectives of
end-point reduction and isoparaffin yield are more effectively
accomplished by processing higher-boiling hydrocarbons; pentanes
and lighter usually comprise less than about 10 mass % of the
naphtha.
The presence of high-boiling compounds is characterized by the end
point, or final boiling point, and/or 95% or 90% distillation point
as measured by the standard ASTM D-86 distillation test; it is well
known that small amounts of high-boiling compounds resulting from
fractionation inefficiencies render final boiling point a less
consistent measure of cut point than 95% or 90% point. Similarly,
the low-boiling cut point may be characterized by 5% or 10% point
rather than initial boiling point. In any event, end points of
reformates are significantly higher than those of the reformer
feeds from which they are derived. The present process combination
enables processing of a naphtha feedstock containing higher-boiling
compounds than otherwise would be possible, according to processes
of the prior art, with high gasoline yields while meeting
reformulated-gasoline specifications. The high-boiling portion of
the naphtha feedstock is converted in the
selective-isoparaffin-synthesis step to obtain a lower-boiling
selective-isoparaffin-synthesis product which can be blended into
gasoline or processed in the reforming zone, thereby converting a
greater proportion of naphtha into gasoline than if a
narrower-range feedstock were processed by catalytic reforming
without selective isoparaffin synthesis.
The naphtha feedstock generally contains small amounts of sulfur
and nitrogen compounds each amounting to less than 10 parts per
million (ppm) on an elemental basis. Preferably the naphtha
feedstock has been prepared from a contaminated feedstock by a
conventional pretreating step such as hydrotreating, hydrorefining
or hydrodesulfurization to convert such contaminants as sulfurous,
nitrogenous and oxygenated compounds to H.sub.2 S, NH.sub.3 and
H.sub.2 O, respectively, which can be separated from hydrocarbons
by fractionation. This conversion preferably will employ a catalyst
known to the art comprising an inorganic oxide support and metals
selected from Groups VIB(6) and VIII(9-10) of the Periodic Table.
[See Cotton and Wilkinson, Advanced Organic Chemistry, John Wiley
& Sons (Fifth Edition, 1988)]. Preferably, the pretreating step
will provide the selective-isoparaffin-synthesis step with a
hydrocarbon feedstock having low sulfur levels disclosed in the
prior art as desirable; e.g., 1 ppm to 0.1 ppm (100 ppb). It is
within the ambit of the present invention that this optional
pretreating step be included in the present process
combination.
A preferred embodiment of the present invention is optimally
understood by reference to FIG. 1. The process combination
comprises a separation zone 10, a reforming zone 20, and an
selective-isoparaffin-synthesis zone 30. Optional units for
light-naphtha isomerization and for dehydrogenation, etherification
or alkylation of synthesis-product isobutane concentrate are not
shown in the Figure, but are discussed hereinafter. For clarity,
only the major sections and interconnections of the process
combination are shown. Individual equipment items such as reactors,
heaters, heat exchangers, separators, fractionators, pumps,
compressors and instruments are well known to the skilled
routineer; description of this equipment is not necessary for an
understanding of the invention or its underlying concepts.
Operating conditions, catalysts, design features and feed and
product relationships are discussed hereinbelow.
In this embodiment, the naphtha feedstock is introduced into
separation zone 10 via line 11. The separation zone generally
comprises one or more fractional distillation columns having
associated appurtenances and separates a heart-cut naphtha fraction
withdrawn via line 12 from a heavy naphtha fraction withdrawn via
line 13. The lower-boiling heart-cut naphtha contains a substantial
concentration of C.sub.7 and C.sub.8 hydrocarbons, which can be
catalytically reformed to produce a reformate component suitable
for blending into current reformulated gasolines. This heart-cut
naphtha also may contain significant concentrations of C6 and C9
hydrocarbons, plus smaller amounts of lower-and higher-boiling
hydrocarbons, depending on the applicable gasoline specifications
and product needs. The heart-cut naphtha end point may range from
about 130.degree. to 175.degree. C., and preferably is within the
range of about 145.degree. to 165.degree. C. The higher-boiling
heavy naphtha contains a substantial amount of C.sub.10
hydrocarbons, and also may contain significant quantities of
lighter and heavier hydrocarbons depending primarily on a petroleum
refiner's overall product balance. The initial boiling point of the
heavy naphtha is between about 120.degree. and 175.degree. C., and
preferably is between 140.degree. and 165.degree. C.
Optionally, a light naphtha fraction may be separated from the
naphtha feedstock in the separation zone via line 14. The light
naphtha comprises pentanes, and may comprise C.sub.6 hydrocarbons.
This fraction is separated from the heart-cut naphtha because
pentanes are not converted efficiently in a reforming zone, and
optionally because C.sub.6 hydrocarbons may be an undesirable feed
to catalytic reforming where they are converted to benzene for
which gasoline restrictions are being implemented. The light
naphtha fraction may be separated from the naphtha feedstock before
it enters the separation zone, in which case this zone would only
separate heart-cut naphtha from heavy naphtha. If the pentane
content of the naphtha feedstock is substantial, however,
separation of light naphtha generally is desirable. This
alternative separation zone generally comprises two fractionation
columns, although in some cases a single column recovering light
naphtha overhead, heavy naphtha from the bottom and heart-cut
naphtha as a sidestream could be suitable.
In this optional embodiment, the heart-cut naphtha fraction is
withdrawn from the separation zone via line 12 and introduced into
reforming zone 20. The reforming zone upgrades the octane number of
the reforming feed through a variety of reactions including
naphthene dehydrogenation and paraffin dehydrocyclization and
isomerization as hereinafter described. It is within the scope of
the invention that the reforming zone also processes heavy
synthesis naphtha from the hereinafter-described
selective-isoparaffin-synthesis zone. Product reformate passes
through line 31 to gasoline blending.
In any event the naphtha feedstock, which may comprise
catalytically cracked gasoline, contains a substantial
concentration of aromatic hydrocarbons, generally ranging from 5 to
80 and more usually 5 to 40 liquid volume percent. These aromatics
may comprise benzene, toluene and higher alkylaromatics within the
boiling ranges described above, and may also comprise small amounts
of naphthalenes and biphenyls within these ranges. The aromatics
generally are not hydrogenated to naphthenes to a large extent in a
naphtha pretreating process as described above, and thus mostly
remain in the feed to a selective isoparaffin-synthesis process of
the prior art. Since aromatics are essentially quantitatively
hydrogenated in a selective isoparaffin-synthesis unit, the
resulting exothermic heat of reaction affects the temperature
profile of the selective isoparaffin-synthesis reaction to a
significant extent.
The present process combination comprises a hydrogenation zone for
saturating aromatic hydrocarbons and a
selective-isoparaffin-synthesis zone. The naphtha feedstock is
charged, along with hydrogen, to the hydrogenation zone which
effects saturation of aromatics at hydrogenation conditions over a
hydrogenation catalyst to produce a saturated intermediate. This
intermediate is transferred to a selective-isoparaffin-synthesis
zone which preferably is contained within the same hydrogen
circuit, i.e., hydrogen and light hydrocarbons are not separated
from the saturated intermediate before the
selective-isoparaffin-synthesis zone. This single circuit obviates
the need for two sets of heat exchangers, separators and
compressors for hydrogen-rich gas. The saturated intermediate thus
also may be transferred to the selective-isoparaffin-synthesis zone
at an increased temperature resulting from the exothermic heat of
the aromatics-hydrogenation reaction. In this manner, heating of
the saturated intermediate optimally is not required. In the
selective-isoparaffin-synthesis zone, the saturated intermediate is
converted to yield lighter products at
selective-isoparaffin-synthesis conditions over a selective
isoparaffin-synthesis catalyst.
Naphtha feedstock and hydrogen comprise combined feed to the
hydrogenation zone. The hydrogenation zone is designed to saturate
aromatics at relatively mild conditions. The hydrogenation zone
contains a bed of catalyst which usually comprises one or more of
nickel and the platinum-group metals, selected from the group
consisting of platinum, palladium, ruthenium, rhodium, osmium, and
iridium, on a suitable refractory inorganic-oxide support. The
inorganic-oxide support preferably comprises alumina, optimally an
anhydrous gamma-alumina with a high degree of purity. The catalyst
advantageously also comprises one or more modifier metals of Groups
VIB (6), VIII (8-10) and IVA (14). Especially preferred catalyst
compositions comprise platinum on an alumina support, treated with
HCI and hydrogen. Alternatively, spent selective
isoparaffin-synthesis catalyst may be used for hydrogenation after
deactivation renders it unsuitable for the synthesis operation.
Such catalysts have been found to provide satisfactory aromatics
saturation at conditions including pressures from about 10 to 100
atmospheres gauge, preferably between about 20 and 70 atmospheres,
and temperatures as low as 30.degree. C. Hydrogen to hydrocarbon
ratios are in the range of about 0.1 to 10, preferably between
about 1 and 5, and liquid hourly space velocities (LHSV) range from
about 1 to 8. In the preferred arrangement of this invention, the
combined feed entering the hydrogenation zone will be heated to a
temperature in the range of 90.degree. to 1200.degree. C. by
indirect heat exchange with the effluent or effluents from the
selective-isoparaffin-synthesis zone. Lower temperatures are found
to be most desirable for the hydrogenation reactions since
nonselective cracking reactions thereby are minimized. Selective
saturation of the aromatics results in a saturated intermediate
from the hydrogenation zone usually containing less than 1 mass %
aromatics. Although hydrogen and light hydrocarbons may be removed
by flash separation and/or fractionation from the saturated
intermediate between the hydrogenation zone and the
selective-isoparaffin-synthesis zone, the intermediate preferably
is transferred between zones without separation of hydrogen or
light hydrocarbons. The exothermic saturation reaction provides a
heated, saturated intermediate to the
selective-isoparaffin-synthesis zone which generally requires no
further heating to effect the required selective
isoparaffin-synthesis temperature. A cooler or other heat exchanger
between the hydrogenation zone and selective-isoparaffin-synthesis
zone may be appropriate for temperature flexibility or for the
startup of the process combination.
Alternative aromatics removal from the naphtha feedstock may be
effected within the scope of the invention by solvent extraction or
adsorptive separation. Solvent extraction for aromatics separation
is well known in the art and may be accomplished using solvent
compositions comprising one or more organic compounds containing at
least one polar group such as a hydroxyl-, amine-, cyano-,
carboxyl-, or nitro- group; preferably the solvent is selected from
one or more of the aliphatic and cyclic alcohols, cyclic monomeric
sulfones, glycols and glycol ethers, glycol esters and glycol ether
esters. Adsorptive separation may be effected using a selective
molecular sieve. This alternative aromatics-removal step features
the advantage of reduced hydrogen consumption and produces an
aromatics concentrate, but does not heat the intermediate sent to
selective isoparaffin synthesis via an exothermic heat of reaction
and reduces the yield of cracked products relative to the preferred
hydrogenation step.
The saturated intermediate has an aromatics content which is
reduced generally about 90% or more relative to the naphtha
feedstock. Usually the aromatics content will be less than about
0.1 mass%, and often in the region of about 100 mass ppm or less
although such low levels are not critical to the utility of the
process combination.
The saturated intermediate is introduced into the
selective-isoparaffin-synthesis zone containing an active,
selective isoparaffin-synthesis catalyst operating at pressures and
temperatures which are significantly below those employed in
conventional hydrocracking. Heavier components of the naphtha are
converted principally to isoparaffins in the presence of hydrogen
with minimum formation of light hydrocarbon gases such as methane
and ethane. Side chains are removed from heavier cyclic compounds
while retaining most of the cyclic rings. Heavy paraffins are
converted to yield a high proportion of isobutane, useful for
production of alkylate or ethers for gasoline blending. The
isobutane generally is at a "superequilibrium" level, i.e., the
proportion of isobutane in total butanes is above the equilibrium
level at synthesis-zone temperatures and thus is higher than could
be achieved by isomerization of butanes. Pentanes formed in the
conversion reaction comprise a high proportion, greater than
generally would be obtained by isomerization, of isopentane, and
other synthesized paraffins also have a preponderance of
branched-chain isomers. The overall effect is that the molecular
weight and final boiling point of the hydrocarbons are reduced,
naphthenic rings are substantially retained, and the content of
isoparaffins is increased significantly in the synthesis product
relative to the naphtha feedstock.
The content of isobutane in total butanes, and usually the content
of isopentane in total pentanes, is greater than the equilibrium
values derived from, e.g., Stull, Daniel L., et al., The Chemical
Thermodynamics of Organic Compounds, 1969, John Wiley and Sons,
esp. pp 245-246 for the relevant operating temperature. This
product ratio is designated herein as superequilibrium isobutane or
isopentane, respectively. Therefore, the present process generally
yields a higher proportion of these isomers than would be
achievable by isomerization of the corresponding paraffin
fraction.
Selective isoparaffin-synthesis operating conditions will vary
according to the characteristics of the feedstock and the product
objectives. Operating pressure may range between about 10
atmospheres and 100 atmospheres gauge, and preferably between about
20 and 70 atmospheres. Temperature is selected to balance
conversion, which is promoted by higher temperatures, against
favorable isomerization equilibrium and product selectivity which
are favored by lower temperatures; operating temperature generally
is between about 50.degree. and 350.degree. C. and preferably
between 100.degree. C. and 300.degree. C. The quantity of catalyst
is sufficient to provide a liquid hourly space velocity of between
about 0.5 and 20, and more usually between about 1.0 an 10. The
operating conditions generally will be sufficient to effect a yield
of at least 8 volume % butanes, and preferably about 15 volume % or
more, from the selective-isoparaffin-synthesis zone relative to the
quantity of saturated intermediate feed to the zone.
Hydrogen is supplied to the reactors of the 'selective
isoparaffin-synthesis process not only to provide for hydrogen
consumed in cracking, saturation and other reactions but also to
maintain catalyst stability. The hydrogen may be partially or
totally supplied from outside the process, but preferably a
substantial proportion of the requirement is provided by hydrogen
recycled after separation from the reactor effluent. The molar
ratio of hydrogen to saturated-intermediate feedstock ranges
usually from about 1.0 to 10, but may be as low as 0.05 to obviate
hydrogen recycle.
The selective-isoparaffin-synthesis zone contains a solid acid
selective isoparaffin-synthesis catalyst. The acid component may
comprise, for example, a halide, such as aluminum chloride, and/or
a zeolite, such as mordenite. The selective isoparaffin-synthesis
catalyst is effective in producing a superequilibrium concentration
of isobutane in butanes produced in the
selective-isoparaffin-synthesis zone at
selective-isoparaffin-synthesis conditions.
The selective isoparaffin-synthesis catalyst preferably comprises
an inorganic-oxide binder, a Friedel-Crafts metal halide and a
Group VIII (8-10) metal component. The refractory inorganic-oxide
support optimally is a porous, adsorptive, high-surface-area
support having a surface area of about 25 to about 500 m.sup.2 /g.
The porous carrier material should also be uniform in composition
and relatively refractory to the conditions utilized in the
process. By the term "uniform in composition," it is meant that the
support be unlayered, has no concentration gradients of the species
inherent to its composition, and is completely homogeneous in
composition. Thus, if the support is a mixture of two or more
refractory materials, the relative amounts of these materials will
be constant and uniform throughout the entire support. It is
intended to include within the scope of the present invention
refractory inorganic oxides such as alumina, titania, zirconia,
chromia, zinc oxide, magnesia, thoria, boria, silica-alumina,
silica-magnesia, chromia-alumina, alumina-boria, silica-zirconia
and other mixtures thereof.
The preferred refractory inorganic oxide for use in the present
invention is alumina. Suitable alumina materials are the
crystalline aluminas known as the gamma-, eta-, and theta-alumina,
with gamma- or eta-alumina giving best results. Zirconia, alone or
in combination with alumina, comprises an alternative
inorganic-oxide component of the catalyst. The preferred refractory
inorganic oxide will have an apparent bulk density of about 0.3 to
about 1.01 g/cc and surface area characteristics such that the
average pore diameter is about 20 to 300 angstroms, the pore volume
is about 0.05 to about 1 cc/g, and the surface area is about 50 to
about 500 m.sup.2 /g.
A particularly preferred alumina is that which has been
characterized in U.S. Pat. Nos. 3,852,190 and 4,012,313 as a
byproduct from a Ziegler higher alcohol synthesis reaction as
described in Ziegler's U.S. Pat. No. 2,892,858. For purposes of
simplification, such an alumina will be hereinafter referred to as
a "Ziegler alumina." Ziegler alumina is presently available from
the Vista Chemical Company under the trademark "Catapal" or from
Condea Chemie GMBH under the trademark "Pural." This material is an
extremely high purity pseudo-boehmite powder which, after
calcination at a high temperature, has been shown to yield a
high-purity gamma-alumina.
The alumina powder may be formed into a suitable catalyst material
according to any of the techniques known to those skilled in the
catalyst-carrier-forming art. Spherical carrier particles may be
formed, for example, from this Ziegler alumina by: (1) converting
the alumina powder into an alumina sol by reaction with a suitable
peptizing acid and water and thereafter dropping a mixture of the
resulting sol and a gelling agent into an oil bath to form
spherical particles of an alumina gel which are easily converted to
a gamma-alumina carrier material by known methods; (2) forming an
extrudate from the powder by established methods and thereafter
rolling the extrudate particles on a spinning disk until spherical
particles are formed which can then be dried and calcined to form
the desired particles of spherical carrier material; and (3)
wetting the powder with a suitable peptizing agent and thereafter
rolling the particles of the powder into spherical masses of the
desired size. This alumina powder can also be formed in any other
desired shape or type of carrier material known to those skilled in
the art such as rods, pills, pellets, tablets, granules,
extrudates, and like forms by methods well known to the
practitioners of the catalyst material forming art.
The preferred form of carrier material for the selective
isoparaffin-synthesis catalyst is a cylindrical extrudate. The
extrudate particle is optimally prepared by mixing the alumina
powder with water and suitable peptizing agents such as nitric
acid, acetic acid, aluminum nitrate, and the like material until an
extrudable dough is formed. The amount of water added to form the
dough is typically sufficient to give a Loss on Ignition (LOI) at
500.degree. C. of about 45 to 65 mass %, with a value of 55 mass %
being especially preferred. The resulting dough is then extruded
through a suitably sized die to form extrudate particles.
The extrudate particles are dried at a temperature of about
150.degree. to about 200.degree. C., and then calcined at a
temperature of about 450.degree. to 800.degree. C. for a period of
0.5 to 10 hours to effect the preferred form of the refractory
inorganic oxide. It is preferred that the refractory inorganic
oxide comprise substantially pure gamma alumina having an apparent
bulk density of about 0.6 to about 1 g/cc and a surface area of
about 150 to 280 m.sup.2 /g (preferably 185 to 235 m.sup.2 /g, at a
pore volume of 0.3 to 0.8 cc/g).
An essential component of the preferred selective
isoparaffin-synthesis catalyst is a platinum-group metal or nickel.
Of the preferred platinum group, i.e., platinum, palladium,
rhodium, ruthenium, osmium and iridium, palladium is a favored
component and platinum is especially preferred. Mixtures of
platinum-group metals also are within the scope of this invention.
This component may exist within the final catalytic composite as a
compound such as an oxide, sulfide, halide, or oxyhalide, in
chemical combination with one or more of the other ingredients of
the composite, or as an elemental metal. Best results are obtained
when substantially all of this metal component is present in the
elemental state. This component may be present in the final
catalyst composite in any amount which is catalytically effective,
and generally will comprise about 0.01 to 2 mass % of the final
catalyst calculated on an elemental basis. Excellent results are
obtained when the catalyst contains from about 0.05 to 1 mass % of
platinum.
The platinum-group metal component may be incorporated into the
selective isoparaffin-synthesis catalyst in any suitable manner
such as coprecipitation or cogellation with the carrier material,
ion exchange or impregnation. Impregnation using water-soluble
compounds of the metal is preferred. Typical platinum-group
compounds which may be employed are chloroplatinic acid, ammonium
chloreplatinate, bromoplatinic acid, platinum dichloride, platinum
tetrachloride hydrate, tetraamine platinum chloride, tetraamine
platinum nitrate, platinum dichlorocarbonyl dichlorfide,
dinitrodiaminoplatinum, palladium chloride, palladium chloride
dihydrate, palladium nitrate, etc. Chloroplatinic acid is preferred
as a source of the especially preferred platinum component.
It is within the scope of the present invention that,the catalyst
may contain other metal components known to modify the effect of
the platinum-group metal component. Such metal modifiers may
include rhenium, tin, germanium, lead, cobalt, nickel, indium,
gallium, zinc, uranium, dysprosium, thallium, and mixtures thereof.
Catalytically effective amounts of such metal modifiers may be
incorporated into the catalyst by any means known in the art.
The composite, before addition of the Friedel-Crafts metal halide,
is dried and calcined. The drying is carried out at a temperature
of about 100.degree. to 300.degree., followed by calcination or
oxidation at a temperature of from about 375.degree. to 600.degree.
C. in an air or oxygen atmosphere for a period of about 0.5 to 10
hours in order to convert the metallic components substantially to
the oxide form.
The resultant oxidized catalytic composite is subjected to a
substantially water-free and hydrocarbon-free reduction step. This
step is designed to selectively reduce the platinum-group component
to the corresponding metal and to insure a finely divided
dispersion of the metal component throughout the carrier material.
Substantially pure and dry hydrogen (i.e., less than 20 vol. ppm
H.sub.2 0) preferably is used as the reducing agent in this step.
The reducing agent is contacted with the oxidized composite at
conditions including a temperature of about 425.degree. C. to about
650.degree. C. and a period of time of about 0.5 to 2 hours to
reduce substantially all of the platinum-group metal component to
its elemental metallic state.
Suitable metal halides comprising the Friedel-Crafts metal
component of the selective isoparaffin-synthesis catalyst include
aluminum chloride, aluminum bromide, ferric chloride, ferric
bromide, zinc chloride and the like compounds, with the aluminum
halides and particularly aluminum chloride ordinarily yielding best
results. Generally, this component can be incorporated into the
catalyst of the present invention by way of the conventional
methods for adding metallic halides of this type; however, best
results are ordinarily obtained when the metallic halide is
sublimed onto the surface of the support according to the preferred
method disclosed in U.S. Pat. No. 2,999,074, which is incorporated
herein by reference.
As aluminum chloride sublimes at about 184.degree. C., suitable
impregnation temperatures range from about 190.degree. C. to
750.degree. C. with a preferable range being from about 500.degree.
C. to 650.degree. C. The sublimation can be conducted at
atmospheric pressure or under increased pressure and in the
presence of absence of diluent gases such a hydrogen or light
paraffinic hydrocarbons or both. The impregnation of the
Friedel-Crafts metal halide may be conducted batch-wise, but a
preferred methodfor impregnating the calcined support is to pass
sublimed AlCl.sub.3 vapors, in admixture With a carrier gas such as
hydrogen, through a calcined catalyst bed. This method both
continuously deposits and reacts the aluminum chloride and also
removes the hydrogen chloride evolved during the reaction.
The amount of Friedel-Crafts metal halide combined with the
calcined support may range from about 1 up to 15 mass % relative to
the calcined composite prior to introduction of the metal-halide
component. The composite containing the sublimed Friedel-Crafts
metal halide is treated to remove the unreacted Friedel-Crafts
metal halide by subjecting the composite to a temperature above the
sublimation temperature of the Friedel-Crafts metal halide,
preferably below about 750.degree. C., for a time sufficient to
remove any unreacted metal halide. In the case of AlCl.sub.3,
temperatures of about 500.degree. C. to 650.degree. C. and times of
from about 1 to 48 hours are preferred.
An optional component of the preferred catalyst is an organic
polyhalo component. In this embodiment, the composite is further
treated preferably after introduction of the Friedel-Crafts metal
halide in contact with a polyhalo compound containing at least 2
chlorine atoms and selected from the group consisting of methylene
halide, haloform, methylhaloform, carbon tetrahalide, sulfur
dihalide, sulfur halide, thionyl halide, and thiocarbonyl
tetrahalide. Suitable polyhalo compounds thus include methylene
chloride, chloroform, methylchloroform, carbon tetrachloride, and
the like. In any case, the polyhalo compound must contain at least
two chlorine atoms attached to the same carbon atom. Carbon
tetrachloride is the preferred polyhalo compound. The composite
contacts the polyhalo compound preferably diluted in a non-reducing
gas such as nitrogen, air, oxygen and the like. The contacting i
suitably is effected at a temperature of from about 100.degree. to
600.degree. C. over a period of from about 0.2 to 5 hours to add at
least 0.1 mass % combined halogen to the composite.
The catalyst of the present invention may contain an additional
halogen component. The halogen component may be either fluorine,
chlorine, bromine or iodine or mixtures thereof with chlorine being
preferred. The halogen component is generally present in a combined
state with the inorganic-oxide support. The halogen component may
be incorporated in the catalyst in any suitable manner, either
during the preparation of the inorganic-oxide support or before,
while or after other catalytic components are incorporated. For
example, chloroplatinic acid may be used in impregnating a platinum
component. The halogen component is preferably well dispersed
throughout the catalyst and may comprise from more than 0.2 to
about 15 mass %, calculated on an elemental basis, of the final
catalyst.
Water and sulfur are catalyst poisons especially for the chlorided
platinum-alumina catalyst composition described hereinabove. Water
can act to permanently deactivate the catalyst by removing
high-activity chloride from this catalyst and replacing it with
inactive aluminum hydroxide. Therefore, water and oxygenates that
can decompose to form water can only be tolerated in very low
concentrations. In general, this requires a limitation of
oxygenates in the feed to about 0.1 ppm or less. Sulfur present in
the feedstock serves to temporarily deactivate the catalyst by
platinum poisoning. If sulfur is present in the feed, activity of
the catalyst may be restored by hot hydrogen stripping of sulfur
from the catalyst composition or by lowering the sulfur
concentration in the incoming feed to below 0.5 ppm. The feed may
be treated by any method that will remove water and sulfur
compounds. Sulfur may be removed from the feed stream by
hydrotreating. Adsorption systems for the removal of sulfur and
water from hydrocarbon streams are well known to those skilled in
the art.
The chlorided platinum-alumina catalyst described hereinabove also
requires the presence of a small amount of an organic chloride
promoter in the selective-isoparaffin-synthesis zone. The organic
chloride promoter serves to maintain a high level of active
chloride on the catalyst, as low levels are continuously stripped
off the catalyst by the hydrocarbon feed. The concentration of
promoter in the combined feed is maintained at from 30 to 300 mass
ppm. The preferred promoter compound is carbon tetrachloride. Other
suitable promoter compounds include oxygen-free decomposable
organic chlorides such as propyldichloride, butylchloride, and
chloroform, to name only a few of such compounds. The need to keep
the reactants dry is reinforced by the presence of the organic
chloride compound which may convert, in part, to hydrogen chloride.
As long as the hydrocarbon feed and hydrogen are dried as described
hereinabove, there will be no adverse effect from the presence of
small amounts of hydrogen chloride.
Contacting within the selective-isoparaffin-synthesis zone may be
effected using the catalyst in a fixed-bed system, a moving-bed
system, a fluidized-bed system, or in a batch-type operation. In
view of the danger of attrition loss of the valuable catalyst and
of operational advantages, it is preferred to use a fixed-bed
system. In this system, a hydrogen-rich gas and the charge stock
are preheated by suitable heating means to the desired reaction
temperature and then passed into a selective-isoparaffin-synthesis
zone containing a fixed bed of the catalyst particle as previously
characterized. The selective-isoparaffin-synthesis zone may be in a
single reactor or in two or more separate reactors with suitable
means therebetween to insure that the desired selective
isoparaffin-synthesis temperature is maintained at the entrance to
each reactor. Two or more reactors in sequence are preferred to
control individual reactor temperatures in light of the exothermic
heat of reaction and for partial catalyst replacement without a
process shutdown. The reactants may be contacted with the bed of
catalyst particles in either upward, downward, or radial flow
fashion. The reactants may be in the liquid phase, a mixed
liquid-vapor phase, or a vapor phase when contacted with the
catalyst particles.
The selective-isoparaffin-synthesis zone generally comprises a
separation section, optimally comprising one or more fractional
distillation columns having associated appurtenances and separating
an isobutane-rich stream, a light synthesis product and a heavy
synthesis product from total synthesis product obtained from the
reaction.
The isobutane-rich stream has a concentration of between about 70
and 95 mole % isobutane in total butanes and more usually in excess
of 80 mole % isobutane. Optionally, an isopentane-rich stream also
may be recovered from the synthesis product either in admixture
with the isobutane or as a separate stream. The isopentane produced
in the selective-isoparaffin-synthesis zone otherwise is recovered
in a light synthesis product fraction which usually is sent to
gasoline blending. The isobutane-rich stream may be further
upgraded via dehydrogenation and etherification or alkylation, as
described hereinafter.
The light synthesis product fraction normally comprises pentanes
and hexanes in admixture, and also may contain smaller
concentrations of naphthenes and C.sub.7 hydrocarbons; benzene
usually is substantially absent. Usually over 80 mole %, and
optimally over 90 mole %, of the C.sub.6 hydrocarbons in the
synthesis product are contained in the light synthesis product;
C.sub.6 hydrocarbons directed to the heavy synthesis product and
subsequently reformed would be partially converted to benzene,
which is undesirable in gasoline for environmental reasons.
In one embodiment, part or all of the isobutane-rich stream is sent
to a dehydrogenation zone. In the dehydrogenation zone, isobutane
is converted selectively to isobutene as feed to etherification
and/or alkylation. Optionally, part or all of the isopentane also
is dehydrogenated to yield isopentene as additional etherification
feed.
A suitable dehydrogenation reaction zone for this invention
preferably comprises one or more radial-flow reactors through which
the catalyst gravitates downward with continuous removal of spent
catalyst, as described in U.S. Pat. No. 3,978,150 which is
incorporated herein by reference. Preferably, the dehydrogenation
reactor section comprises multiple stacked or side-by-side
reactors, and a combined stream of hydrogen and hydrocarbons is
processed serially through the multiple reactors each of which
contains a particulate catalyst disposed as an annular-form moving
bed. The moving catalyst bed permits continuous addition of fresh
and/or regenerated catalyst and the withdrawal of spent catalyst,
and is illustrated in U.S. Pat. No. 3,647,680 which is incorporated
by reference. Since the dehydrogenation reaction is endothermic in
nature, intermediate heating of the reactant stream between
reactors is the optimal practice.
Dehydrogenation conditions generally include a pressure of from
about 0 to 35 atmospheres, more usually no more than about 5
atmospheres. Suitable temperatures range from about 480.degree. C.
to 760.degree. C., optimally from about 540.degree. C. to
705.degree. C. when processing a light liquid comprising isobutane
and/or isopentane. Hydrogen is admixed with the hydrocarbon
feedstock in a mole ratio of from about 0.1 to 10, and more usually
from about 0.5 to 2. Catalyst is available in dehydrogenation
reactors to provide a liquid hourly space velocity of from about 1
to 10, and preferably no more than about 5.
The dehydrogenation catalyst comprises a platinum-group metal
component, preferably a platinum component, and an alkali-metal
component on a refractory support. The alkali-metal component is
chosen from cesium, rubidium, potassium, sodium, and lithium. The
catalyst also may contain promoter metals, preferably tin in an
atomic ratio of tin to platinum be between 1:1 and about 6:1. The
refractory support of the dehydrogenation catalyst should be a
porous, absorptive high-surface-area material as delimited
hereinabove for the reforming catalyst. A refractory inorganic
oxide is the preferred support, with alumina being particularly
preferred.
The dehydrogenation zone will produce a near-equilibrium mixture of
the desired isoolefin and its isoalkane precursor. Preferably an
isobutane-rich stream is processed to yield an isobutene-containing
stream. Alternatively or additionally, an isopentene-containing
stream is produced from and isopentane-rich stream. A separation
section recovers hydrogen from the product for use elsewhere.
Optionally part or all of an olefin-containing product stream from
the dehydrogenation zone is used to produce ethers in an
etherification zone. The olefin-containing stream preferably
contains isobutene, and may comprise isopentene. In addition, one
or more monohydroxy alcohols are fed to the etherification zone.
Ethanol is a preferred monohydroxyoalcohol feed, and methanol is
especially preferred. This variety of possible feed materials
allows the production of a variety of ethers in addition to or
instead of the preferred methyl tertiary-butyl ether (MTBE). These
useful ethers include ethyl tertiary butyl ether (ETBE), methyl
tertiary amyl ether (MTAE) and ethyl tertiary amyl ether
(ETAE).
Processes operating with vapor, liquid or mixed-phase conditions
may be suitably employed in this invention. The preferred
etherification process uses liquid-phase etherification conditions,
including a superatmospheric pressure sufficient to maintain the
reactants in liquid phase but no more than about 50 atmospheres;
even in the presence of additional light materials, pressures in
the range of 10 to 40 atmospheres generally are sufficient to
maintain liquid-phase conditions. Operating temperature is between
about 30.degree. C. and 100.degree. C.; the reaction rate is
normally faster at higher temperatures, but conversion is more
complete at lower temperatures. High conversion in a moderate
volume reaction zone can, therefore, be obtained if the initial
section of the reaction zone, e.g., the first two-thirds, is
maintained above 70.degree. C. and the remainder of the reaction
zone is maintained below 50.degree. C. This may be accomplished
most easily with two reactors.
The ratio of feed alcohol to isoolefin should normally be
maintained in the broad range of 1:1 to 2:1. With the preferred
reactants good results are achieved if the ratio of methanol to
isobutene is between 1.05:1 and 1.5:1. An excess of methanol, above
that required to achieve satisfactory conversion at good
selectivity, should be avoided as some decomposition of methanol to
dimethylether may occur with a concomitant increase in the load on
separation facilities.
A wide range of materials are known to be effective as
etherification catalysts including mineral acids such as sulfuric
acid, boron trifluoride, phosphoric acid on kieselguhr,
phosphorus-modified zeolites, heteropoly acids, and various
sulfonated resins. The use of a sulfonated solid resin catalyst is
preferred. These resin type catalysts include the reaction products
of phenolformaldehyde resins and sulfuric acid and sulfonated
polystyrene resins including those cross-linked with
divinylbenzene. Further information on suitable etherification
catalysts may be obtained by reference to U.S. Pat. Nos. 2,480,940,
2,922,822, and 4,270,929 and the previously cited etherification
references.
In the preferred etherification process for the production of MTBE,
essentially all of the isobutene is converted to MTBE thereby
eliminating the need for subsequently separating that olefin from
isobutane. As a result, downstream separation facilities are
simplified. Several suitable etherification processes have been
described in the literature which presently are being used to
produce MTBE. The preferred form of the etherification zone is
similar to that described in U.S. Pat. No. 4,219,678. In this
instance, the isobutene, methanol and a recycle stream containing
recovered excess alcohol are passed into the etherification zone
and contacted at etherification conditions with an acidic
etherification catalyst to produce an effluent containing MTBE.
The effluent from the etherification-zone reactor section includes
at least product ethers, light hydrocarbons, dehydrogenatable
hydrocarbons, and any excess alcohol. The effluent may also include
small amounts of hydrogen and of other oxygen-containing compounds
such as dimethyl ether and TBA. The effluent passes from the
etherification reactor section to a separation section for the
recovery of product. The etherification effluent is separated to
recover the ether product, preferably by fractional distillation
with ether being taken as bottoms product; this product generally
is suitable for gasoline blending but may be purified further,
e.g., by azeotropic distillation.
The overhead from ether separation containing unreacted
hydrocarbons is passed through a methanol recovery zone for the
recovery of methanol, preferably by adsorption, with return of the
methanol to the etherification reactor section. The
hydrocarbon-rich stream is fractionated to remove C.sub.3 and
lighter hydrocarbons and oxygenates from the stream of unreacted
C.sub.4 -C.sub.5 hydrocarbons. Heavier oxygenate compounds are
removed by passing the stream of unreacted hydrocarbons through a
separate oxygenate recovery unit. This hydrocarbon raffinate, after
oxygenate removal, may be dehydrogenated to provide additional
feedstock for the etherification zone or used as part of the feed
to an alkylation reaction zone to produce high octane alkylate.
A portion of the isobutane-rich stream from the separation section
and a portion of the iso-olefin-containing stream from the
dehydrogenation zone may be processed in an alkylation zone. The
alkylation zone optionally may process other isobutane- or
olefin-containing streams from an associated petroleum
refinery.
The optional alkylation zone of this invention may be any acidic
catalyst reaction system such as a hydrogen fluoride-catalyzed
system, sulfuric-acid system or one which utilizes an acidic
catalyst in a fixed-bed reaction system. Hydrogen fluoride
alkylation is particularly preferred, and may be conducted
substantially as set forth in U.S. Pat. No. 3,249,650. The
alkylation reaction in the presence of hydrogen fluoride catalyst
is conducted at a catalyst to hydrocarbon volume ration within the
alkylation reaction zone of from about 0.2 to 2.5 and preferably
about 0.5 to 1.5. Ordinarily, anhydrous hydrogen fluoride will be
charged to the alkylation system as fresh catalyst; however, it is
possible to utilize hydrogen fluoride containing as much as 10.0%
water or more. Excessive dilution with water is generally to be
avoided since it tends to reduce the alkylating activity of the
catalyst and further introduces corrosion problems. In order to
reduce the tendency of the olefinic portion of the charge stock to
undergo polymerization prior to alkylation, the molar proportion of
isoparaffins to olefinic hydrocarbons in an alkylation reactor is
desirably maintained at a value greater than 1.0, and preferably
from about 3.0 to 15.0. Alkylation reaction conditions, as
catalyzed by hydrogen fluoride, include a temperature of from
-20.degree. to about 100.degree. C., and preferably from about
0.degree. to 50.degree. C. The pressure maintained within the
alkylation system is ordinarily at a level sufficient to maintain
the hydrocarbons and catalyst in a substantially liquid phase; that
is, from about atmospheric to 40 atmospheres. The contact time
within the alkylation reaction zone is conveniently expressed in
terms of space-time, being defined as the volume of catalyst within
the reactor contact zone divided by the volume rate per minute of
hydrocarbon reactants charged to the zone. Usually the space-time
will be less than 30 minutes and preferably less than about 15
minutes.
Alkylate recovered from the alkylation zone generally comprises
n-butane and heavier components, with isobutane and lighter,
materials having been removed by fractionation and returned to the
reactor. At least a portion, and preferably all, of the alkylate is
blended into gasoline.
It is within the scope of the invention that a portion of the light
synthesis product, especially the C.sub.6 portion, is isomerized in
an isomerization zone. Usually, the C.sub.5 portion would not be
upgraded by isomerization, since the pentanes already generally
comprise an isopentane/n-pentane ratio in excess of equilibrium at
usual isomerization conditions.
Contacting within the isomerization zone may be effected using the
catalyst in a fixed-bed system, a moving-bed system, a
fluidized-bed system, or in a batch-type operation. A fixed-bed
system is preferred. The isomerization zone may be in a single
reactor or in two or more separate reactors with suitable means
therebetween to insure that the desired isomerization temperature
is maintained at the entrance to each zone. Two or more reactors in
sequence are preferred to enable improved isomerization through
control of individual reactor temperatures and for partial catalyst
replacement without a process shutdown. The reactants may be
contacted with the bed of catalyst particles in either upward,
downward, or radial-flow fashion. The reactants may be in the
liquid phase, a mixed liquid-vapor phase, or a vapor phase when
contacted with the catalyst particles, with excellent results being
obtained by application of the present invention to a primarily
liquid-phase operation.
Isomerization conditions in the isomerization zone include reactor
temperatures usually ranging from about 400.degree. to 250.degree.
C. Lower reaction temperatures are generally preferred in order to
favor equilibrium mixtures having the highest concentration of
high-octane highly branched isoalkanes and to minimize cracking of
the feed to lighter hydrocarbons. Temperatures in the range of from
about 40.degree. to about 150.degree. C. are preferred in the
present invention. Reactor operating pressures generally range from
about atmospheric to 100 atmospheres, with preferred pressures in
the range of from 20 to 35 atmospheres. Liquid hourly space
velocities range from about 0.25 to about 12 volumes of
isomerizable hydrocarbon feed per hour per volume of catalyst, with
a range of about 0.5 to 5 hr.sup.-1 preferred.
Hydrogen is admixed with the feed to the isomerization zone to
provide a mole ratio of hydrogen to hydrocarbon feed of about 0.0.1
to 5. The hydrogen may be supplied totally from outside the process
or supplemented by hydrogen recycled to the feed after separation
from reactor effluent. Light hydrocarbons and small amounts of
inerts such as nitrogen and argon may be present in the hydrogen.
Water should be removed from hydrogen supplied from outside the
process, preferably by an adsorption system as is known in the art.
In a preferred embodiment the hydrogen to hydrocarbon mol ratio in
the reactor effluent is equal to or less than 0.05, generally
obviating the need to recycle hydrogen from the reactor effluent to
the feed.
Any catalyst known in the art to be suitable for the isomerization
of paraffin-rich hydrocarbon streams may be used as an
isomerization catalyst in the isomerization zone. One suitable
isomerization catalyst comprises a platinum-group metal,
hydrogen-form crystalline aluminosilicate and a refractory
inorganic oxide, and the composition preferably has a surface area
of at least 580 m.sup.2 /g. The preferred noble metal is platinum
which is present in an amount of from about 0.01 to 5 mass % of the
composition, and optimally from about 0.15 to 0.5 mass %.
Catalytically effective amounts of one or more promoter metals
preferably selected from Groups VIB(6), VIII(8-10), IB(11),
IIB(12), IVA(14), rhenium, iron, cobalt, nickel, gallium and indium
also may be present. The crystalline aluminosilicate may be
synthetic or naturally occurring, and preferably is selected from
the group consisting of FAU, LTL, MAZ and MOR with mordenite having
a silica-to-alumina ratio of from 16:1 to 60:1 being especially
preferred. The crystalline aluminosilicate generally comprises from
about 50 to 99.5 mass % of the composition, with the balance being
the refractory inorganic oxide. Alumina, and preferably one or more
of gamma-alumina and eta-alumina, is the preferred inorganic oxide.
Further details of the composition are disclosed in U.S. Pat. No.
4,735,929, incorporated herein by reference thereto.
A preferred isomerization catalyst composition comprises one or
more platinum-group metals, a halogen, and an inorganic-oxide
binder. Preferably the catalyst contains a Friedel-Crafts metal
halide, with aluminum chloride being especially preferred. The
optimal platinum-group metal is platinum which is present in an
amount of from about 0.1 to 0.5 mass %. The composition may also
contain an organic polyhalo component, with carbon tetrachloride
being preferred, and the total chloride content is from about 2 to
10 mass %. The inorganic oxide preferably comprises alumina, with
one or more of gamma-alumina and eta-alumina providing best
results. Optimally, the carrier material is in the form of a
calcined cylindrical extrudate. Other details and alternatives of
preparation steps and operation of the preferred isomerization
catalyst are as presented hereinabove for the selective
isoparaffin-synthesis catalyst. Optionally, the same catalyst may
be used in the selective isoparaffin-synthesis and isomerization
zones. U.S. Pat. Nos. 2,999,074 and 3,031,419 teach additional
aspects of this composition and are incorporated herein by
reference.
Isomerate recovered from once-through processing of light naphtha
does contain some low-octane normal paraffins and
intermediate-octane methylhexanes as well as the desired
highest-octane isopentane and dimethylbutane. It is within the
scope of the present invention that the product from the reactors
of the isomerization process is subjected to separation and recycle
of the lower-octane portion to the isomerization reaction.
Low-octane normal paraffins are separated and recycled in this
embodiment to obtain an iso-rich product, and less-branched hexanes
also may be separated and recycled. Techniques to achieve this
separation are well known in the art, and include fractionation and
molecular-sieve adsorption.
The heart-cut naphtha from naphtha separation and/or heavy
synthesis product optionally may be processed in a reforming zone
to obtain a reformate product of increased octane number. Reforming
may be carried out in two or more fixed-bed reactors in sequence or
in moving-bed reactors with continuous catalyst regeneration.
Reforming operating conditions include a pressure of from about
atmospheric to 60 atmospheres (absolute), with the preferred range
being from atmospheric to 20 atmospheres and a pressure of below 10
atmospheres being especially preferred. Hydrogen is supplied to the
reforming zone in an amount sufficient to correspond to a ratio of
from about 0.1 to 10 moles of hydrogen per mole of hydrocarbon
feedstock. The operating temperature generally is in the range of
260.degree. to 560.degree. C. The volume of the contained reforming
catalyst corresponds to a liquid hourly space velocity of from
about 1 to 40 hr.sup.-1.
The reforming catalyst conveniently is a dual-function composite
containing a metallic hydrogenation-dehydrogenation component on a
refractory support which provides acid sites for cracking,
isomerization, and cyclization. The hydrogenation-dehydrogenation
component comprises a supported platinum-group metal component,
with a platinum component being preferred. The platinum may exist
within the catalyst as a compound, in chemical combination with one
or more other ingredients of the catalytic composite, or as an
elemental metal; best results are obtained when substantially all
of the platinum exists in the catalytic composite in a reduced
state. The catalyst may contain other metal components known to
modify the effect of the preferred platinum component, including
Group IVA (14) metals, other Group VIII (8-10) metals, rhenium,
indium, gallium, zinc, uranium, dysprosium, thallium and mixtures
thereof with a tin component being preferred.
The refractory support of the reforming catalyst should be a
porous, adsorptive, high-surface-area material which is uniform in
composition. Preferably the support comprises refractory inorganic
oxides such as alumina, silica, titania, magnesia, zirconia,
chromia, thoria, boria or mixtures thereof, especially alumina with
gamma- or eta-alumina being particularly preferred and best results
being obtained with "Ziegler alumina" as described in the
references. Optional ingredients are crystalline zeolitic
aluminosilicates, either naturally occurring or synthetically
prepared such as FAU, MEL, MFI, MOR, MTW (IUPAC Commission on
Zeolite Nomenclature), and non-zeolitic molecular sieves such as
the aluminophosphates of U.S. Pat. No. 4,310,440 or the
silico-aluminophosphates of U.S. Pat. No. 4,440,871 (incorporated
by reference). Further details of the preparation and activation of
embodiments of the above reforming catalyst are disclosed in U.S.
Pat. No. 4,677,094 (Moser et al.), which is incorporated into this
specification by reference thereto.
In an advantageous alternative embodiment, the reforming catalyst
comprises a large-pore molecular sieve. The term "large-pore
molecular sieve" is defined as a molecular sieve having an
effective pore diameter of about 7 angstroms or larger. Examples of
large-pore molecular sieves which might be incorporated into the
present catalyst include LTL, FAU, AFI, MAZ, and zeolite-beta, with
a nonacidic L-zeolite (LTL) being especially preferred. An
alkali-metal component, preferably comprising potassium, and a
platinum-group metal component, preferably comprising platinum, are
essential constituents of the alternative reforming catalyst. The
alkali metal optimally will occupy essentially all of the cationic
exchangeable sites of the nonacidic L-zeolite. Further details of
the preparation and activation of embodiments of the alternative
reforming catalyst are disclosed, e.g., in U.S. Pat. Nos. 4,619,906
(Lambert et al) and 4,822,762 (Ellig et al.), which are
incorporated into this specification by reference thereto.
Preferably part or all of each of the synthesis product and
optional light synthesis product, ether, alkylate, isomerized
product and reformate are blended to produce a gasoline component.
Finished gasoline may be produced by blending the gasoline
component with other constituents including but not limited to one
or more of butanes, butenes, pentanes, naphtha, catalytic
reformate, somerate, alkylate, polymer, aromatic extract, heavy
aromatics; gasoline from catalytic cracking, hydrocracking, thermal
cracking, thermal reforming, steam pyrolysis add coking; oxygenates
from sources outside the combination such as methanol, ethariol,
propanol, isopropanol, TBA, SBA, MTBE, ETBE, MTAE and higher
alcohols and ethers; and small amounts of additives to promote
gasoline stability and uniformity, avoid corrosion and weather
problems, maintain a clean engine and improve driveability.
EXAMPLES
The following examples serve to illustrate certain specific
embodiments of the present invention. These examples should not,
however, be construed as limiting the scope of the invention as set
forth in the claims. There are many possible other variations, as
those of ordinary skill in the art will recognize, which are within
the spirit of the invention.
Example 1
The feedstock used in Examples 2 and 3 is a full-range naphtha
derived from a paraffinic mid-continent crude oil and has the
following characteristics:
______________________________________ Specific gravity 0.746
Distillation, ASTM D-86, .degree.C. IBP 86 50% 134 EP 194 Mass %
paraffins 63.7 naphthenes 24.0 aromatics 12.3 Mass % C.sub.5 0.2
C.sub.6 12.1 C.sub.7 19.3 C.sub.8 20.2 C.sub.9 + 48.2
______________________________________
Example 2
The benefits of using the process combination of the invention are
illustrated by contrasting results with those from a corresponding
process of the prior art. Example 2 presents results based on the
use of a .prior-art process combination.
The prior art is illustrated by selective isoparaffin synthesis
from the naphtha feedstock described above without prior
hydrogenation of the aromatics in the feedstock. A pilot plant was
loaded with (i) quartz chips and (ii) a platinum-AlCl.sub.3
-on-alumina selective isoparaffin-synthesis catalyst as described
hereinabove in a volumetric ratio of (i):(ii) of 4:5. The quartz
chips served for effective control of the temperature of the feed
to the selective-isoparaffin-synthesis zone. The selective
isoparaffin-synthesis catalyst contained about 0.25 mass % platinum
and 5.5 mass % chloride.
Selective isoparaffin synthesis from the naphtha feedstock was
effected at a pressure of about 30 atmospheres and a
hydrogen-to-hydrocarbon mol ratio of 2.5. Tests were carried out at
inlet temperatures of 120.degree., 150.degree. and 180.degree. C. A
temperature profile was constructed by measuring temperatures at 20
points across the catalyst bed. The profile is shown in FIG. 2.
In order to assess the impact of the invention on isoparaffin
selectivity, ratios of isobutane/total butanes and isopentane/total
pentenes were measured for the prior-art operation. These
isoparaffin/total-paraffin ratios are shown, along with feed
conversion to pentanes and lighter products, in FIG. 3.
Example 3
Results from applying the process combination of the invention are
illustrated in Example 3. Selective isoparaffin synthesis from the
naphtha feedstock described above was effected following
hydrogenation of the aromatics in the feedstock. A pilot plant was
loaded with (a) a chlorided platinum-alumina catalyst, (b) quartz
chips and (c) a platinum-AlCl.sub.3 -on-alumina selective
isoparaffin-synthesis catalyst as described hereinabove in a
volumetric ratio of (a): (b): (c) of 5:7:15. As in Example 2, the
selective isoparaffin-synthesis catalyst contained about 0.25 mass
% platinum and 5.5 mass % chloride.
The combination of aromatics saturation and selective isoparaffin
synthesis from the naphtha feedstock was effected at a pressure of
about 30 atmospheres and a hydrogen-to-hydrocarbon mol ratio of
2.5. Tests were carried out at inlet temperatures of 120.degree.,
150.degree. and 1800.degree. C. A temperature profile was
constructed by measuring temperatures at 20 points across the
catalyst bed. The profile is contrasted with that of the prior art
in FIG. 2. Note that peak temperature across the selective
isoparaffin-synthesis bed is less than 10.degree. C. above inlet
temperature, in contrast to the prior art for which the temperature
increase is in the range of about 25.degree. to 40.degree. C.
In order to assess the impact of the invent ion on isoparaffin
selectivity, product ratios of isobutane/total butaries and
isopentane/total pentanes were measured for the processes of the
invention and of the prior art. These isoparaffin/total-paraffin
ratios are shown in FIG. 3. Note that the isobutane/butane, ratios
of the invention are somewhat higher than those of the prior art
and the isopentane/pentane ratios are substantially higher when
using the process combination of the invention. Considering
equilibrium values of isobutane/total butanes as derived from Stull
et al., supra, the proportion of isobutane is 0.26 to 0.28 higher
than equilibrium and therefore the butane product contains
superequilibrium isobutane.
Example 4
The feedstock used in Examples 5-11 is a mixture of heavy straight
naphtha and coker naphtha derived from Arabian Light crude oil and
having the following characteristics:
______________________________________ Specific gravity 0.758
Distillation, ASTM D-86, .degree.C. IBP 93 50% 137 90% 168 EP 197
Volume % paraffins 63.3 naphthenes 19.2 aromatics 17.5 Volume %
C.sub.6 - 0.4 C.sub.7 20.8 C.sub.8 27.0 C.sub.9 26.8 C.sub.10 16.7
C.sub.11 + 8.3 ______________________________________
Example 5
The benefits of producing a gasoline component using the process
combination of the invention are illustrated by contrasting results
with those from processes of the prior art. Example 5 presents
results based on the use of a prior-art process combination.
The prior art is illustrated by selective isoparaffin synthesis
from the naphtha feedstock described above followed by
fractionation of the effluent and reforming of the C.sub.7 and
heavier synthesis naphtha. Yields in the synthesis zone are based
on the use of a platinum-AlCl.sub.3 -on-alumina catalyst as
described hereinabove containing about 0.25 mass % platinum and 5.5
mass % chloride. Hydrogen consumption and product yields based on
the processing of 3250 cubic meters per day are as follows:
______________________________________ Hydrogen consumption,
10.sup.3 Nm .sup.3 /day 642 Yields, m.sup.3 /day: Isobutane
concentrate 1339 C.sub.5 /C.sub.6 1098 C.sub.7 + 1321
______________________________________
The isobutane concentrate ("Iso C.sub.4 concentrate") comprises
about 90% isobutane.
C.sub.7 + product from selective isoparaffin synthesis is processed
in a reforming unit. The reforming operation is carried out using
each of two alternative catalyst types:
Case A: Conventional spherical platinum-tin-alumina
Case B: Platinum on potassium-form L-zeolite extrudate
Operating pressure in each case is about 3.4 atmospheres gauge, and
the severity is 95 Research octane number (RON) clear on the
C.sub.5 + product. The low pressure provides high hydrogen and
C.sub.5 + yields. After stabilization of the reformate to remove
the small amount of C.sub.4 and lighter produced, the C.sub.5 + is
blended with C.sub.5 /C.sub.6 from selective isoparaffin synthesis
to obtain a gasoline component. Overall yields of the
selective-isoparaffin-synthesis/reforming combination, considering
hydrogen production in the reformer as well as consumption in the
selective isoparaffin synthesis, are as follows:
______________________________________ Case: A B
______________________________________ Net H.sub.2 consumption,
10.sup.3 Nm .sup.3 /day 267 203 Yields, m.sup.3 /day: Iso C.sub.4
concentrate 1339 1339 C.sub.5 + component 2176 2163
______________________________________
Example 6
The process combination of the invention is illustrated in Example
6. The Arabian Light naphtha as used in control Example 5 is
fractionated to separate a 150.degree. C. and heavier cut from a
cut boiling up to about 150.degree. C. in accordance with the
Figure. The heavier naphtha is processed by selective isoparaffin
synthesis followed by fractionation of the effluent to yield an
isobutane concentrate, a C.sub.5 /C.sub.6 fraction and heavy
synthesis naphtha. Yields in the selective-isoparaffin-synthesis
zone are based on the use of a platinum-AlCl.sub.3 -on-alumina
catalyst as described hereinabove containing about 0.25 mass %
platinum and 5.5 mass % chloride. Hydrogen consumption and product
yields based on the processing of 3250 cubic meters per day are as
follows:
______________________________________ IBP-150.degree. C. naphtha
to reforming, m.sup.3 /day 1552 150.degree. C. and heavier naphtha
to synthesis, m.sup.3 /day 1698 Hydrogen consumption, 10.sup.3 Nm
.sup.3 /day 359 Synthesis, yields, m.sup.3 /day: IsoC.sub.4
concentrate 704 C.sub.5 /C.sub.6 1056 C.sub.7 + 230
______________________________________
The isobutane concentrate comprises about 90% isobutane.
The C.sub.7 + synthesis naphtha is processed along with
IBP-150.degree. C. naphtha in a reforming unit. The reforming
operation is carried out using an extruded catalyst comprising
platinum on potassium-form L-zeolite at a pressure of about 3.4
atmospheres gauge and a severity of 95 Research octane number (RON)
clear on the C.sub.5 + product. After stabilization of the
reformats to remove the small amount of C.sub.4 and lighter
produced, the C.sub.5 + is blended with C.sub.5 /C.sub.6 from
selective isoparaffin synthesis to obtain a gasoline component.
Overall yields of the selective-isoparaffin-synthesis/reforming
combination, considering hydrogen production in the reformer i as
well as consumption in selective isoparaffin synthesis, are as
follows from 3250 cubic meters/day of naphtha:
______________________________________ Net H.sub.2 production,
10.sup.3 Nm .sup.3 /day 143 Yields, m.sup.3 /day: IsoC.sub.4
concentrate 704 C.sub.5 + component 2540
______________________________________
Compared to Example 5 of the prior art, the, corresponding case of
the invention shows lower isobutane production but a greater
C.sub.5 + yield and net production rather than consumption of
hydrogen.
Example 7
Example 7 presents a reforming process Of the prior art producing a
gasoline component which has an unacceptable endpoint for current
U.S. reformulated gasoline blends. The feedstock to the reforming
process is the same Arabian Light naphtha used in Example 5. The
reforming operation is carried out using a conventional spherical
platinum-rhenium-on-alumina catalyst at a pressure of about 20
atmospheres gauge and a severity of 92 Research octane number (RON)
Clear on the C.sub.5 + product.
Yields of hydrogen and C.sub.5 + reformate and high-end
distillation characteristics of the reformate are as follows:
______________________________________ Net H.sub.2 production,
10.sup.3 Nm .sup.3 /day 386 Yields, m.sup.3 /day: C.sub.5 +
component 2766 C.sub.5 + ASTM D-86: 90% point, .degree.C. 173 End
point, .degree.C. 214 ______________________________________
Example 8
Example 8 is another illustration of the prior ,art based on the
selective isoparaffin synthesis of the naphtha feedstock described
above followed by fractionation of the effluent and reforming of
the C.sub.7 and heavier synthesis naphtha using operating
conditions in accordance with Example 7. Yields in the
selective-isoparaffin-synthesis zone are identical to those of
Example 5.
C.sub.7 + product from selective isoparaffin synthesis is processed
in a reforming unit. As in Example 7, the reforming operation is
carried out using a conventional spherical
platinum-rhenium-on-alumina catalyst at a pressure of about 20
atmospheres gauge and a severity of 92 Research octane number (RON)
clear on the C.sub.5 + product. After stabilization of the
reformats to remove the small amount of C.sub.4 and lighter
produced, the C.sub.5 + is blended with C.sub.5 /C.sub.6 from
selective isoparaffin synthesis to obtain a gasoline component.
Overall yields of the selective-isoparaffin-synthesis/reforming
combination, considering hydrogen production in the reformer as
well as consumption in the selective isoparaffin synthesis and the
high end distillation characteristics of the reformate, are as
follows:
______________________________________ Net H.sub.2 consumption,
10.sup.3 Nm .sup.3 /day 361 Yields, m.sup.3 /day: IsoC.sub.4
concentrate 1339 C.sub.5 + component 2184 C.sub.5 + ASTM D-86: 90%
point, .degree.C. 139 End point, .degree.C. 165
______________________________________
Example 9
Example 9 is an illustration of the process combination of the
invention for comparison with prior-art Examples 7 and 8. The
Arabian Light naphtha described hereinabove is fractionated to
separate a 175.degree. C. and heavier cut from a cut boiling up to
about 175.degree. C. in accordance with the Figure. The lighter cut
contains 26 volume % C.sub.7 and 34 volume % C.sub.8 hydrocarbons.
The 175.degree. C. and heavier cut contains about 58 volume %
C.sub.10 hydrocarbons. The heavier naphtha is processed by
selective isoparaffin synthesis followed by fractionation of the
effluent to yield isobutane concentrate and a C.sub.5 + synthesis
product. Yields in the selective-isoparaffin-synthesis zone are
based on the use of a platinum-AlCl.sub.3 -on-alumina catalyst as
described hereinabove containing about 0.25 mass % platinum and 5.5
mass % chloride. Product yields from fractionation and synthesis
are as follows:
______________________________________ IBP - 175.degree. C. naphtha
to reforming, m.sup.3 /day 2604 175.degree. C. and heavier naphtha
to synthesis, m.sup.3 /day 646 Hydrogen consumption, 10.sup.3 Nm
.sup.3 /day 120 Yields, m.sup.3 /day: IsoC.sub.4 concentrate 226
C.sub.5 + synthesis product 537
______________________________________
The isobutane concentrate comprises about 90% isobutane. The
IBP-175.degree. C. naphtha is processed in a reforming unit using a
conventional spherical platinum-rhenium-on-alumina catalyst as in
Example 3 at an operating pressure of about 20 atmospheres gauge
and a severity of 92 Research octane number (RON) clear on the
C.sub.5 + product. After stabilization of the reformate to remove
the small amount of C.sub.4 and lighter produced, the C.sub.5 + is
blended with C.sub.5 /C.sub.6 from selective isoparaffin synthesis
to obtain a gasoline component. Overall yields of the
selective-isoparaffinsynthesis/reforming combination, considering
hydrogen production in the reformer as well as consumption in the
selective isoparaffin synthesis, are as follows from 3250 cubic
meters/day of naphtha:
______________________________________ Net H.sub.2 production,
10.sup.3 Nm .sup.3 /day 243 Yields, m.sup.3 /day: IsoC.sub.4
concentrate 226 C.sub.5 + component 2788 C.sub.5 + ASTM D-86: 90%
point, .degree.C. 166 End point, .degree.C. 193
______________________________________
Example 10
Example 10 is another illustration of the process combination of
the invention, based on a change in cut point between light and
heavy naphtha. The Arabian Light naphtha described hereinabove is
fractionated to separate a 160.degree. C. and heavier cut from a
cut boiling up to about 160.degree. C. in accordance with the
Figure. The lighter cut contains 34 volume % C.sub.7 and 44 volume
% C.sub.8 hydrocarbons. The 160.degree. C. and heavier cut contains
about 43 volume % C.sub.10 hydrocarbons. The heavier naphtha is
processed by selective isoparaffin synthesis followed by
fractionation of the effluent to yield products according to two
different cases:
Case A: ISOC.sub.4 concentrate and C.sub.5 + synthesis product
Case B: ISOC.sub.4 concentrate, C.sub.5 /C.sub.6 fraction, Heavy
synthesis naphtha
Thus, the C.sub.5 + product in Case B is separated into a C.sub.5
/C.sub.6 cut to gasoline blending and a C.sub.7 + fraction as
reforming feed.
Yields in the selective-isoparaffin-synthesis zone are based on the
use of a platinum-AlCl.sub.3 -on-alumina catalyst as described
hereinabove containing about 0.25 mass % platinum and 5.5 mass %
chloride. Product yields from fractionation and synthesis are as
follows:
______________________________________ IBP - 160.degree. C. naphtha
to reforming, m.sup.3 /day 1991 160.degree. C. and heavier naphtha
to synthesis, m.sup.3 /day 1259 Hydrogen consumption, 10.sup.3 Nm
.sup.3 /day 260 Yields, m.sup.3 /day: IsoC.sub.4 concentrate 456
C.sub.5 /C.sub.6 753 C.sub.7 + 270
______________________________________
The isobutane concentrate comprises about 90% isobutane.
In Case A, the entire C.sub.5 + effluent from selective isoparaffin
synthesis is blended into the gasoline component and reforming feed
consists of the Arabian Light naphtha cut boiling up to about
160.degree. C. In Case B, the C.sub.7 + synthesis naphtha is added
to the feed to the reforming unit. The reforming operation is
carried out using a conventional spherical
platinum-rhenium-on-alumina catalyst as in Example 7 at a pressure
of about 20 atmospheres gauge and a severity of 92 Research octane
number (RON) clear on the C.sub.5 + product. After stabilization of
the reformate to remove the small amount of C.sub.4 and lighter
produced, the C.sub.5 + is blended with C.sub.5 /C.sub.6 from
selective isoparaffin synthesis to obtain a gasoline component.
Overall yields of the selective-isoparaffin-synthesis/reforming
combination from 3250 cubic meters per day of naphtha, considering
hydrogen production in the reformer as well as consumption in the
selective isoparaffin synthesis, are as follows:
______________________________________ Case: A B
______________________________________ Net H.sub.2 production,
10.sup.3 Nm.sup.3 /day -21 75 Yields, m.sup.3 /day: IsoC.sub.4
concentrate 456 456 C.sub.5 + component 2695 2692 C.sub.5 + ASTM
D-86: 90% point, .degree.C. 144 148 End point, .degree.C. 176 177
______________________________________
Example 11
A comparison of the cases of Examples 7-10 shows the impact on
yields of using the present invention to reduce the end point of a
gasoline component, based on 3250 cubic meters per day of naphtha
feed:
______________________________________ Example 7 8 9 10 Case A B
______________________________________ Invention? No No Yes Yes Yes
Synthesis feed, m.sup.3 /day 0 3250 646 1259 1259 Net H.sub.2,
10.sup.3 Nm .sup.3 /day 386 -361 243 -21 75 IsoC.sub.4 .div.
C.sub.5 +, m.sup.3 /day 2766 3523 3014 3151 3148 C.sub.5 +, m.sup.3
/day 2766 2184 2788 2695 2692 90% point, .degree.C. 173 139 166 144
148 End point, .degree.C. 214 165 193 176 177
______________________________________
The invention enables end-point reduction with very little C.sub.5
+ loss and some gain in C.sub.4 +. The net hydrogen production is
reduced with the addition of selective isoparaffin synthesis, but a
favorable balance may be maintained in the
reforming/selective-isoparaffin-synthesis combination.
The process combination of the invention is surprisingly effective
in increasing the yield of isoparaffins from a selective
isoparaffin-synthesis process, thus providing higher product
octanes and more potential for valuable isoparaffin derivatives
such as ethers and alkylate.
Example 12
The feedstock used in Examples 13-14 is a catalytically cracked
gasoline from which most of the C.sub.5 and C.sub.6 have been
removed and which has the following characteristics:
______________________________________ Specific gravity 0.815
Distillation, ASTM D-86, .degree.C. IBP 62 50% 166 90% 216 EP 233
Volume % paraffins 18.9 olefins 15.7 naphthenes 14.3 aromatics 51.1
______________________________________
Example 13
The catalytically cracked gasoline of Example 12 was hydrotreated
severely to eliminate the effect of sulfur and nitrogen on
subsequent tests. The hydrotreating was carried out at 52
atmospheres in three successive beds of catalyst comprising
cobalt-molybdenum on alumina to achieve the following contaminant
levels:
______________________________________ Sulfur, mass ppm 0.3
Nitrogen, mass ppm <0.1
______________________________________
Subsequent tests of selective isoparaffin synthesis with 2 ppm
sulfur added to the feed showed an average difference in
temperature profile through the synthesis reactor of only about
1.degree. C., indicating that the depth of hydrotreating required
is not as substantial as originally believed.
The hydrotreated product was fractionated to yield a feedstock to
selective isoparaffin synthesis having the following
characteristics:
______________________________________ Specific gravity 0.808
Distillation, ASTM D-86, .degree.C. IBP 110 5% 125 10% 131 50% 164
90% 209 95% 219 EP 239 ______________________________________
Example 14
The benefits of using the process combination of the invention are
illustrated by contrasting results with those from a corresponding
process of the prior art. The comparison is carried out in a
similar manner to that presented in Examples 2 and 3.
The prior art is illustrated by selective isoparaffin synthesis
from the hydrotreated catalytically cracked gasoline feedstock
described above without prior hydrogenation of the aromatics in the
feedstock. A pilot plant was loaded with (i) quartz chips and (ii)
a platinum-AlCl.sub.3 -on-alumina selective isoparaffin-synthesis
catalyst as described hereinabove in a volumetric ratio of (i):(ii)
of about 7:15. The quartz chips served for effective control of the
temperature of the feed to the selective-isoparaffin-synthesis
zone. The selective isoparaffin-synthesis catalyst contained about
0.25 mass % platinum and 5.5 mass % chloride.
The process combination of the invention was illustrated by an
alternative pilot-plant configuration to hydrogenate aromatics in
the feedstock prior to selective isoparaffin synthesis. A pilot
plant was loaded with (a) a chlorided platinum-alumina catalyst,
(b) quartz chips and (c) a platinum-Alcl.sub.3 -on-alumina
selective-isoparaffin-synthesis catalyst as described hereinabove
in a ratio of (a):(b):(c) of about 5:9:18.
Selective isoparaffin synthesis from the naphtha feedstock was
effected in each of the above reactor loadings at a pressure of
about 30 atmospheres, a hydrogen-to-hydrocarbon mol ratio of 2.5,
and an inlet temperature of about 200.degree. C. A temperature
profile was constructed by measuring temperatures at 20 points
across the catalyst bed. The profiles of the comparative runs are
shown in FIG. 4.
* * * * *